Aromatics Và Các Phương Pháp Sản Xuất

10.6.1 Aromatics productionand use

IntroductionAn aromatics complex is a combination of

process units that can be used to convertpetroleum naphtha and pyrolysis gasoline (pygas)into the basic petrochemical intermediates:Benzene Toluene and Xylenes (BTX). Benzene isa versatile petrochemical building block used inthe production of more than 250 differentproducts. Ethylbenzene, cumene, andcyclohexane are the most important benzenederivatives. The xylenes product, also known asmixed xylenes, contains four different aromaticisomers: para-xylene, ortho-xylene, meta-xylene,and ethylbenzene. Small amounts of mixedxylenes are used for solvent applications, butmost xylenes are processed further within thecomplex to produce one or more of the individualisomers. The most important C8 aromatic isomeris para-xylene, which is used almost exclusivelyfor the production of polyester fibres, resins, andfilms. In recent years, polyester fibres haveshown growth rates of 5 to 6 percent per year andresin has shown growth rates of 10 to 15 percentper year, corresponding to the emergence of PET(polyethylene terephthalate) containers. A smallamount of toluene is recovered for use in solventapplications and derivatives, but most toluene isused to produce benzene and xylenes. Toluene isbecoming increasingly important for theproduction of xylenes through toluenedisproportionation and transalkylation with C9aromatics.

Aromatics complexes can have manydifferent configurations. The simplest complexproduces only benzene, toluene, and mixed

xylenes and consists of the following majorprocess units: naphtha hydrotreating for theremoval of sulphur and nitrogen contaminants;catalytic reforming for the production ofaromatics from naphtha; aromatics extractionfor the extraction of BTX.

However, most modern aromatics complexesare designed to maximize the yield of benzeneand para-xylene and sometimes ortho-xylene.About one-half of the existing UOP (UniversalOil Products) aromatics complexes areconfigured for the production of bothpara-xylene and ortho-xylene. These neweraromatics complexes have not only the majorprocess units mentioned above, but includethe following additional major process units:para-xylene extraction for the separationof para-xylene for the other xylene isomers;xylene isomerization for production of anequilibrium mixture of xylene isomers; tolueneand C9 aromatic transalkylation for productionof xylenes and benzene.

An aromatics complex may be configured inmany different ways, depending on the availablefeedstocks, the desired products, and the amountof investment capital available. Because of thiswide flexibility, the product slate can be varied tomatch downstream processing requirements.

Feedstock considerationsAny of the following streams may be used as

feedstock to an aromatics complex: straight-runnaphtha; hydrocracked naphtha; mixed xylenes;pyrolysis gasoline; coke-oven light oil;condensate; Liquid Petroleum Gas (LPG).

Petroleum naphtha is by far the most popularfeedstock for aromatics production. Reformednaphtha, or reformate, accounts for 70% of total

591VOLUME II / REFINING AND PETROCHEMICALS

10.6

Aromatics

world BTX supply. The pygas by-product fromethylene plants is the next-largest source at 23%.Coal liquids from coke ovens account for theremaining 7%. Pygas and coal liquids areimportant sources of benzene that may be usedonly for benzene production or may be combinedwith reformate and fed to an integrated aromaticscomplex. Mixed xylenes are also actively tradedand can be used to feed a stand-aloneParex-Isomar loop or to provide supplementalfeedstock for an integrated complex.

CCR Platforming processThe CCR (Continuous Catalytic Reformer)

Platforming process is used throughout the worldtoday in the petroleum and petrochemicalindustries. It produces aromatics fromnaphthenes and paraffins, either for use in motorfuel or as a source of specific aromaticcompounds. In aromatics applications, the feednaphtha is generally restricted to C6 through C9compounds to maximize the production ofbenzene, toluene, and xylenes.

The distribution of the hydrocarbon types(paraffins, naphthenes and aromatics) willdetermine how easily various naphthas can bereformed. Aromatic compounds pass through thePlatforming unit relatively unchanged.Naphthenes react rapidly and efficiently toaromatics, while paraffins react slowly and withless selectivity.

Process chemistryFour major reactions occur in the reactors to

produce the desired products: dehydrocyclizationof paraffins to 5-membered rings; isomerizationof 5-membered to 6-membered rings;

dehydrogenation of 6-membered rings toaromatics; hydrocracking of large hydrocarbonsto smaller hydrocarbons. The function of thereformer is to efficiently convert paraffins andnaphthenes to aromatics with as little ringopening or cracking as possible.

Process descriptionHydrotreated naphtha feed is combined with

recycle hydrogen gas and heat exchanged againstreactor effluent (Fig. 1). The combined feed isthen raised to reaction temperature in the chargeheater and sent to the reactor section. Typicallyfour adiabatic radial-flow reactors are arrangedin a vertical stack. The catalyst flows verticallyby gravity down the stack, while the charge flowsradially across the annular catalyst beds. Thepredominant reactions are endothermic, so aninterheater is used between each reactor to reheatthe charge to reaction temperature. Flue gas fromthe fired heaters is typically used to generatehigh pressure steam, but other heat integrationoptions are available.

The effluent from the last reactor is heatexchanged against combined feed, cooled andsplit into vapour and liquid products in aseparator. The vapour phase is rich in hydrogengas. A portion of the gas is compressed andrecycled back to the reactors. The net hydrogen-rich gas is compressed and charged together withthe separator liquid phase to the product recoverysection. This section can be engineered andoptimized to provide the required performance.The liquid product from the recovery section issent to a stabilizer where light saturates areremoved from the aromatics-rich reformateproduct. Over time, coke builds up on the

592 ENCYCLOPAEDIA OF HYDROCARBONS

BULK PRODUCTS AND PRODUCTION LINES IN THE PETROCHEMICAL INDUSTRY

combinedfeed

exchanger

naphtha feedfrom treating

net gascompressor

net H2 rich gas

fuel gas

light ends

aromaticsrich reformate

stabilizer

sepa

rato

r

CCRregenerator

regeneratedcatalyst

spentcatalyst

fired heaters

stackedreactor

Fig. 1. CCR Platforming process.

Platforming catalyst at reaction conditions. Apartially deactivated catalyst is continuallywithdrawn from the bottom of the reactor stackand transferred to the CCR regenerator. Thecatalyst flows down through the regeneratorwhere the accumulated carbon is burned off andthe moisture and chloride levels are adjusted.Regenerated catalyst is lifted with hydrogen tothe top of the reactor stack. Because the reactorand regenerator sections are separate, each canoperate at its own optimum conditions. Inaddition, the regenerator section can betemporarily shut down for maintenance withoutaffecting the operation of the reactor and productrecovery sections.

Process performanceAn understanding of process chemistry

explains the loss of volumetric liquid yield acrossthe reaction zone. Yield loss comes from twosources: natural shrinkage resulting from thehigher density of aromatic and cracking reactionsthat form lower value light products.

The conversion of naphthenes and paraffins toaromatics causes an increase in the density of thereaction material. Fig. 2 shows the volumetricconversion of typical lean and rich naphthasacross the Platforming process. A lean naphtha isone in which the paraffin content of the feed ishigh, typically above 65%. Rich naphtha has alower paraffin content and a smaller density shiftbetween feed and reformate. The catalyst systemcannot control the volumetric change in yieldscaused by the density increase to aromatics.However, by decreasing hydrocracking reactions,selective catalyst systems have a tremendouseffect on yield loss. Coupling a selective catalystsystem with optimized reaction conditions

provides an increase in the catalytic reactionselectivity and the suppression of thehydrocracking reactions. Both of which are key to attaining maximum aromatic and hydrogenyields.

The fact that yields in the reforming reactionsystem are favoured by low pressure is wellknown. Great advances have been made over thepast two decades in coupling improved catalystcompositions with lower operating pressures. Ateffective reactor operating pressures of 3.5�105

Pa (50 psi), which are typical of UOP’s latestdesigns and recent commercial start-ups, theselectivities of the more difficult reactions areclearly better. Under such conditions, reactionselectivities for heavier paraffin species andheavy 5-membered and 6-membered naphthenering species range from 80 to 100%. Thus,through pressure reduction and using currentcatalyst technology, dramatic progress has beenachieved towards closing the actual-to-theoreticalyield gap.

The lower operating pressures increase therate of coke formation on the catalyst and cancause an eventual loss of performance. Thisproblem was solved in 1971 with the start-up ofthe world’s first CCR regenerator, which wasdesigned and developed by UOP. In the yearssince the start-up of this first unit, UOP hascontinued to improve and expand regenerationcapability to keep the CCR regenerator system inbalance with the requirements of the reactorsection.

Cyclar processThe Cyclar process converts LPG directly

into a liquid aromatics product in a singleoperation. UOP, working jointly with anothercompany, developed the Cyclar processexpanding the use of LPG to the production ofhigh-value petrochemical aromatics. LPGconsists mainly of the propane and butanefraction recovered from gas and oil fields andpetroleum refining operations. The relatively lowvalue and abundance of LPG make it an idealfeedstock for petrochemical applications.Benzene, toluene, and xylenes are producedprimarily through the catalytic reforming ofpetroleum naphtha. However, naphtha is in greatdemand for gasoline and petrochemicalproduction and the value of naphtha is expectedto rise as supplies become tighter. The Cyclarprocess offers a unique ability to producepetrochemical-grade BTX from a lower-valuefeedstock, and can be used in production fields


AROMATICS

leannaphtha

reformate

from PN

P

N

P

A

A

loss

from N

from A

richnaphtha

reformate

from PN

P

N

P

A

A

loss

from N

from A

Fig. 2. Platforming process reactions; P, paraffins; N, naphthenes; A, aromatics.

to convert excess LPG into a liquid product forpipeline transport.

Process chemistryThe Cyclar process converts LPG directly to a

liquid aromatics product in a single operation.The reaction is best described asdehydrocyclodimerization, and isthermodynamically favoured at temperatureshigher than 425°C. The dehydrogenation of lightparaffins (propane and butanes) to olefins is therate limiting step. Once formed, the highlyreactive olefins oligomerize to form largerintermediates, which then rapidly cyclize tonaphthenes. These reactions, dehydrogenation,oligomerization, and cyclization, are all acidcatalyzed.

The final reaction step is the dehydrogenationof the naphthenes to their correspondingaromatics. This reaction is highly favoured atCyclar process operating conditions, resulting invirtually complete conversion of the naphthenes.The intermediates can also undergo ahydrocracking side reaction to form methane andethane, resulting in a yield loss. Heavierparaffins such as pentanes can also be includedin the feed. Olefins and pentanes are almostcompletely converted in the Cyclar unit, but theunit must be designed to handle them, since theyresult in a higher catalyst coking rate than purebutane and propane feedstocks. Although thereaction sequence involves exothermic steps, thepreponderance of dehydrogenation reactionsresult in a highly endothermic overall reaction.This is easily understood from the fact that fivemoles of hydrogen are produced for every moleof aromatics formed from propane or butane.

Process descriptionThe Cyclar process is divided into three major

sections (Fig. 3). The reactor section includes aradial-flow reactor stack, combined feedexchanger, charge heater, and interheaters.

The CCR regenerator section includes aregenerator stack and catalyst transfer system.The product recovery section includes productseparators, compressors, stripper, and gasrecovery equipment. The flow scheme is similarto that of the CCR Platforming process, which isused widely throughout the world for reformingpetroleum naphtha. Fresh feed and recycle arecombined and heat exchanged against reactoreffluent. The combined feed is then raised toreaction temperature in the charge heater andsent to the reactor section. Four adiabatic,

radial-flow reactors are arranged in one or morevertical stacks. Catalyst flows by gravity downthe stack, while the charge flows radially acrossthe annular catalyst beds. Between each reactor,the charge is reheated to reaction temperature inan interheater. The effluent from the last reactoris split into vapour and liquid products in aseparator. The liquid is sent to a stripper wherelight saturates are removed from the C6�

aromatic product. Vapour from the separator iscompressed and sent to a gas recovery section,typically a cryogenic unit, for separation into a95% pure hydrogen product stream, a fuel gasstream of light saturates, and a recycle stream ofunconverted LPG.

Over time, coke builds up on the Cyclarcatalyst at reaction conditions. The partiallydeactivated catalyst is continually withdrawn fromthe bottom of the reactor stack and transferred tothe CCR regenerator for regeneration. The catalystflows down through the regenerator where theaccumulated carbon is burned off. Regeneratedcatalyst is lifted with hydrogen to the top of thereactor stack. Because the reactor and regeneratorsections are separate, each can operate at its ownoptimum conditions. In addition, the regeneratorsection can be temporarily shut down formaintenance without affecting the operation of thereactor and product recovery sections.

The principal Cyclar operating variables aretemperature, space velocity, pressure, andfeedstock composition. The temperature must behigh enough to ensure nearly completeconversion of reaction intermediates in order toproduce a liquid product that is essentially freeof non-aromatic impurities, but low enough tominimize nonselective thermal reactions. Spacevelocity is optimized against conversion withinthis temperature range to obtain high productyields with minimum operating costs. Reactionpressure has a big impact on processperformance. UOP currently offers twoalternative Cyclar process designs. The lowpressure design is recommended wheremaximum aromatic yield is desired. The high-pressure design requires only half the catalyst and is attractive where minimuminvestment and operating costs are the overridingconsiderations.

Product quality and yieldsThe major liquid products from the Cyclar

process are benzene, toluene, xylenes, andheavier aromatics. In general, the aromatics yieldincreases with the carbon number of the



feedstock. In a low-pressure operation, theoverall aromatics yield increases from 61 wt% offresh feed with an all-propane feedstock to 66%with an all-butane feed, with a correspondingdecrease in fuel gas production. These yieldfigures can be interpolated linearly for mixedpropane and butane feedstocks. The distributionof butane species in the feed has no effect onyields. The distribution of aromatic species in theliquid product is also unaffected by feedstockcomposition. Butane feedstocks produce aproduct that is leaner in benzene and richer inxylenes than that produced from propane. Witheither propane or butane feeds, the liquid productcontains about 91% BTX and 9% heavieraromatics.

The Cyclar process produces high-qualityaromatic products. Petrochemical grade tolueneand xylenes can be obtained by fractionationalone, without the need for subsequentextraction. The by-product light ends containsubstantial amounts of hydrogen, which may berecovered in several different ways, depending onthe purity desired. An absorber/stripper systemproduces a 65 mol% hydrogen product stream; acold box produces 95 mol% hydrogen; anabsorber-stripper system combined with aPressure Swing Adsorption (PSA) unit produces99 mol% hydrogen; and a cold box combinedwith a PSA unit can produce 99 � mol%hydrogen, if desired.

RZ Platforming process RZ Platforming process is a fixed bed system

that is well suited for use in aromatics productionfacilities, particularly for those producers whorequire large amounts of benzene. The RZPlatforming process uses the RZ-100 catalyst. Byvirtue of its ability to convert the most difficultfeed components (C6 and C7 paraffins) toaromatics, the RZ-100 catalyst represents a majorstep beyond conventional reforming catalysttechnology.

The RZ Platforming process is primarily usedfor situations where higher yields of BT(Benzene Toluene) aromatics and hydrogen aredesired. Benzene production especially benefitsfrom RZ-100’s catalyst selectivity performance.The superior ability of the RZ Platformingprocess to handle light, paraffinic feeds and itsflexibility in processing straight-run naphthafractions provide many options for improvingaromatics production and supplying neededhydrogen, either in new units or in existingaromatics facilities.

Process chemistryThe Platforming process is designed to

efficiently convert paraffins and naphthenes toaromatics with as little ring opening or crackingas possible.

Although RZ-100 catalyst is similar in manyways to conventional reforming catalysts, it


AROMATICS

combinedfeed

exchanger

fresh LPG feedfrom driers

light endsdrag to fuel gas

C6+ aromatic product

hydrogen product fuel gas

separator

stripper

CCRregenerator

regeneratedcatalyst

spentcatalyst

fired heaters

stackedreactor

Fig. 3. Cyclar process.

differs greatly in the production of lightaromatics – benzene and toluene. The selectivityof conventional reforming catalysts for benzeneand toluene is significantly lower than for the C8aromatics. By comparison, the selectivity of theRZ-100 catalyst for light aromatics is vastlyimproved.

Process descriptionThe RZ Platforming unit configuration is

similar to other fixed bed platforming units (Fig. 4). Treated naphtha feed is combined withrecycle hydrogen gas and heat exchanged againstreactor effluent. The combined feed is then raisedto reaction temperature in the charge heater andsent to the reactor section. Radial-flow reactorsare arranged in a conventional side-by-sidepattern. The predominant reactions areendothermic, so an interheater is used betweeneach reactor to reheat the charge to reactiontemperature.

Flue gas from the fired heaters is typicallyused to generate high pressure steam, but otherheat integration options are available. Theeffluent from the last reactor is heat exchangedagainst combined feed, cooled and split intovapour and liquid products in a separator. Thevapour phase is rich in hydrogen gas and aportion of the gas is compressed and recycledback to the reactors. The net hydrogen-rich gas iscompressed and charged together with theseparator liquid phase to the product recoverysection. This section can be engineered andoptimized to provide location-specific requiredperformance. The liquid product from therecovery section is sent to a stabilizer where lightsaturates are removed from the C6� aromatic

product. Catalysts deactivate over time atreaction conditions. Typical cycle lengths areeight to twelve months. Efficient ex situregeneration facilities for rejuvenation of RZ-100catalysts are available.

Process performanceAlthough the CCR Platforming process is the

most efficient means possible for producingxylenes from heavier naphtha fractions, itsconversion of C6 and C7 paraffins to aromatics isnormally below 50%, even at low pressure. TheRZ Platforming process offers constant aromaticsselectivity, in the range of 80% or higher, evenwhen processing the most difficult C6 and C7paraffin feed components.

Feedstock to the RZ Platforming unit canrange from extraction-unit raffinate to BTXnaphtha. A very effective application for theRZ-100 catalyst is the production of aromaticsand hydrogen from light, paraffinic feeds, suchas a BT raffinate. The RZ-100 catalyst can alsobe used in parallel with a conventional reformingunit to optimize the production of the desiredaromatics by processing different fractions of thehydrotreated feed. In such cases, the conventionalreformer can be dedicated to process the heavierfeed fraction, taking advantage of its superiorability to produce xylenes. The light naphtha,which is rich in C6 and C7 components, can berouted to the RZ Platforming unit, whereselectivity for converting light paraffins tobenzene and toluene is greatest.

Parex processThe Parex process is an innovative, adsorptive

separation method for the recovery of



combinedfeed

exchanger

naphtha feedfrom treating

net gascompressor

net H2 rich gas

fuel gas

light ends

stabilizer

sepa

rato

r

fired heaters

reactors

C6+ aromatics

Fig. 4. RZ Platforming process.

para-xylene from mixed xylenes that offers highproduct purity, high product recovery, highon-stream efficiency, and extended adsorbentlife. ‘Mixed xylenes’ refers to a mixture of C8aromatic isomers that includes ethylbenzene,para-xylene, meta-xylene, and ortho-xylene.These isomers boil so closely together thatseparating them by conventional distillation isnot practical. The Parex process provides anefficient means of recovering para-xylene using asolid zeolitic adsorbent which is selective forpara-xylene. Unlike conventional batchchromatography, the Parex process simulates amoving bed of adsorbent with continuouscounter-current flow of a liquid feed over theadsorbent. Feed and products enter and leave theadsorbent bed continuously, at nearly constantcompositions. This technique is calledsimulating-moving-bed separation.

In an aromatics complex, the Parex unit islocated downstream of the xylenes column and isintegrated with an Isomar unit. The feed to thexylenes column consists of the C8� aromaticsproduct from a CCR Platforming unit and thexylenes produced in a Tatoray unit. The C8fraction from the overhead of the xylenes columnis fed to the Parex unit, where high purity para-xylene is recovered in the extract. The Parexunit raffinate is sent to the Isomar unit, where theother C8 aromatic isomers are converted intoadditional para-xylene and recycled back to thexylene column. The Parex process is designed torecover more than 97 wt% of the para-xylenefrom the feed in a single pass while deliveringproduct purity of 99.9 wt% or better.

The high purity para-xylene recovered in theParex process is used for the production ofpolyester fibres, resins, and films. Para-xylene isconverted to terephthalic acid (TPA) or dimethylterephthalate (DMT), which are then reacted withethylene glycol to form polyethyleneterephthalate (PET) which is the raw material formost polyesters.

Process description The separation takes place in the adsorbent

chambers. Each adsorbent chamber is dividedinto a number of adsorbent beds. Each bed ofadsorbent is supported from below by specializedinternals or grids, that are designed to producehighly efficient flow distribution. Each internalsassembly is connected to the rotary valve by abed line. The internals between each adsorbentbed are used to inject or withdraw liquid from thechamber while simultaneously collecting liquid

from the bed above and redistributing the liquidover the bed below. A typical Parex unit has 24adsorbent beds with 26 sets of chamber internals,and 24 bed lines connecting the beds to therotary valve. Due to practical constructionconsiderations, most Parex units consist of twoadsorption chambers in series with 12 beds each.In the Parex process, there are four major streamsthat are distributed to the adsorbent chamber bythe rotary valve. These ‘net’ streams include: a) feed in (mixed xylenes feed); b) dilute extractout (para-xylene product diluted with desorbent);c) dilute raffinate out (ethylbenzene, meta- andortho-xylene diluted with desorbent); d ) desorbent in (recycle desorbent from thefractionation section).

At any given time, only four of the bed linesare active, carrying the net streams into and outof the adsorbent chamber. The rotary valve isused to periodically switch the positions of theliquid feed and withdrawal points as thecomposition profile moves down the chamber. Apump provides the liquid circulation from thebottom of the first adsorbent chamber to the topof the second. A second pump providescirculation from the bottom of the secondadsorbent chamber to the top of the first. In thisway, the two adsorbent chambers function as asingle continuous loop of adsorbent beds. Thedilute extract is sent to the extract column forseparation of the extract from the desorbent. Theoverhead from the extract column is sent to afinishing column where the highly pure para-xylene product is separated from anytoluene which may have been present in the feed.

The dilute raffinate from the rotary valve issent to the raffinate column for separation of theraffinate from the desorbent. The overhead fromthe raffinate column contains unextracted C8aromatic components: ethylbenzene, meta-xylene, ortho-xylene, and any non-aromatics which may have been present inthe feed. The raffinate product is then sent to theIsomar unit where additional para-xylene isformed and then recycled back to the Parex unit.

The desorbent from the bottom of both theextract and raffinate columns is recycled back tothe adsorbent chambers through the rotary valve.In order to prevent this accumulation, provisionis made to take a slipstream of the recycledesorbent to a small desorbent rerun columnwhere any heavy contaminants are rejected.During normal operation, mixed xylenes arestripped, clay-treated, and rerun prior to beingsent to the Parex unit. Therefore, the amount of


AROMATICS

heavy contaminants to be removed from thebottom of the desorbent rerun column is notsignificant.

Isomar processThe Isomar process is used to maximize the

recovery of a particular xylene isomer from amixture of C8 aromatic isomers. The Isomarprocess is most often applied to para-xylenerecovery, but it can also be used to maximize therecovery of ortho-xylene or meta-xylene. In thecase of para-xylene recovery, a mixed xylenesfeed is charged to a Parex process unit where thepara-xylene isomer is preferentially extracted.The raffinate from the Parex unit, almost entirelydepleted of para-xylene, is then sent to theIsomar unit. The Isomar unit re-establishes anequilibrium distribution of xylene isomers,essentially creating additional para-xylene fromthe remaining ortho- and meta-xylenes. Effluentfrom the Isomar unit is then recycled back to theParex unit for recovery of additional para-xylene. In this way, the ortho- and meta-xylenes and ethylbenzene are recycled toextinction. Depending on the type of catalyst,ethylbenzene is converted into xylenes orbenzene.

Process chemistryThere are two broad categories of xylene

isomerization catalysts: ethylbenzene (EB)isomerization catalysts, which convertethylbenzene into additional mixed xylenes, andethylbenzene dealkylation catalysts, whichconvert ethylbenzene to a valuable benzeneco-product.

UOP currently offers I-400 EB isomerizationcatalyst and I-300 EB dealkylation catalyst. Theselection of the isomerization catalyst depends onthe configuration of the aromatics complex, thecomposition of the feedstocks, and the desiredproduct slate. The I-400 catalyst is typicallychosen when the primary goal of the complex isto maximize production of para-xylene.Alternatively, I-300 catalyst can be used to de-bottleneck an existing Parex unit orcrystallizer by converting more EB per passthrough the isomerization unit and eliminatingthe requirement for naphthene intermediatecirculation around the Parex-Isomar recycle loop.The EB isomerization reaction presents athermodynamic limit, insofar as the EBconversion to equilibrium is equal to about 30wt% per pass. The EB dealkylation reaction isnot thermodynamic limited, allowing for up to 70

wt% or greater EB conversion per pass. Thereduction in the size of the Parex-Isomar loopwith an I-300 reload comes at the expense oflower paraxylene yields, since all the EB in thefeed is being converted to benzene rather thanadditional para-xylene. All xylene isomerizationcatalysts cause some loss of aromatic rings to by-products across the reactor. A large portion ofthe total feed from the Isomar unit goes to thexylene column. A typical Parex-Isomar loopexhibits a combined feed ratio of about 3.5.Therefore, a small reduction in the per-pass by-product formation across the Isomar unittranslates to a large yield advantage. The Isomarprocess exhibits minimal by-product formation,with the precise level being a function of feedcomposition, catalyst type and operating severity.For EB isomerization catalysts, the by-productformation is based on the per-pass losses of C8cyclics (EB, xylenes and C8 naphthenes). For theI-400 catalyst, the ring loss per pass will be inthe range of 1.5-2 mol% per pass. The C8 ringloss for the I-9 catalyst ranges from 3 to 5 mol%per pass. For EB dealkylation catalysts, the by-product formation is based on the per-passlosses of xylenes. For I-300 catalyst, this is inthe range of 1-2 wt%. The EB is converted tobenzene with a selectivity that is typically morethan 90 mol%. Overall aromatic ringconservation is very high – more than 99 mol%for EB dealkylation catalysts. Most of the by-products from Isomar (except for the benzenefrom I-300) catalyst can be re-converted toxylenes when a Tatory transalkylation unit ispart of the flow-scheme.

I-300 catalyst offers the simplicity of a singlecatalyst system. The Isomar process requiresabout half as much I-300 catalyst as the previousgeneration of catalysts, and its use eliminatesmultiple beds of different catalysts withcomplicated loadings and distributors. Unlikesome ethylbenzene dealkylation catalysts, the I-300 catalyst does not require continuousaddition of ammonia to achieve desired activityand selectivity.

Process descriptionAn Isomar unit is always combined with a

recovery unit for one or more of the xyleneisomers. Usually it is combined with a Parex unitfor recovery of para-xylene. In the Parex-Isomarflow scheme (Fig. 5), fresh mixed xylenes are fedto the xylene column, which can be designedeither to recover ortho-xylene in the bottoms orsimply reject C9� aromatic components in order



to meet feed specifications to the Parex unit. Thexylene column overhead is then directed to theParex unit where 99.9 wt% para-xylene isproduced. The raffinate from the Parex unit,containing less than 1 wt% para-xylene, is sent tothe Isomar unit.

Process performanceThe best way to compare xylene

isomerization catalysts is to measure the overallpara-xylene yield from the Parex-Isomar loop.The para-xylene yield, based on fresh mixed xylenes fed to the Parex-Isomar loop, is characterized by the following considerations.

The basis for the comparison is the Parex-Isomar yield flow scheme, processing afresh mixed xylenes feed consisting of 17 wt%ethylbenzene, 18 wt% para-xylene, 40 wt% meta-xylene, and 25 wt% ortho-xylene. Theoperating severity for the I-9 and I-400 catalyst is 22.1 wt% para-xylene of the totalxylenes from the Isomar unit. The operatingseverity for I-300 catalyst is 65 wt% conversionof ethylbenzene per pass. With the I-9 catalyst,the overall yield of para-xylene is 84 wt% of thefresh mixed xylenes feed. Since it has very high overall aromatic ring retention, butconverts ethylbenzene to benzene, the I-300catalyst exhibits a higher overall yield of benzeneplus para-xylene, but the yield of para-xylene isonly 76.5 wt%. This means that more mixedxylenes are required to produce a target amount

of para-xylene with I-300 catalyst. I-400 catalystrelies on the same reaction chemistry as I-9catalysts, but is more selective and exhibits lower ring loss. With I-400 catalyst, the overallyield of para-xylene is 7 wt% higher than with I-9 catalysts, at 91wt% of fresh mixedxylenes feed.

Tatoray process The Tatoray process is used to selectively

convert toluene and C9 aromatics (A9) intobenzene and xylenes. In a modern aromaticscomplex, this process is integrated between thearomatics extraction and xylene recoverysections of the plant. Extracted toluene is fed tothe Tatoray process unit rather than beingblended into the gasoline pool or sold for solventapplications.

To maximize the production of para-xylenefrom the complex, the A9 by-product can also befed to the Tatoray process unit. This shifts thechemical equilibrium from benzene production toxylenes production. In recent years, the demandfor para-xylene has outstripped the supply ofmixed xylenes. The Tatoray process provides anideal way to produce additional mixed xylenesfrom toluene and heavy aromatics. Incorporatinga Tatoray process unit into an aromatics complexcan more than double the yield of para-xylenefrom a given naphtha feedstock.

Process chemistryThe two major reactions in the Tatoray

process are disproportionation andtransalkylation. The conversion of toluene intobenzene and xylenes is called toluenedisproportionation. Transalkylation is theconversion of a mixture of toluene and A9 intoxylenes.

This process is designed to function at a muchhigher level of conversion per pass than othertoluene disproportionation processes. With atypical 50:50 feedstock ratio of toluene and C9aromatics, the overall conversion isapproximately 50% per pass. This highconversion level minimizes the amount ofunconverted material that must be recycled backthrough the BT fractionation section of thecomplex. A smaller recycle stream minimizes thesize of the benzene and toluene columns, the sizeof the Tatoray process unit, and the utilityconsumption of all of these units.

The Tatoray process reactions are conductedin a hydrogen atmosphere to minimize cokeformation on the catalyst. Because there is


AROMATICS

chargeheater

purgegas

make-uphydrogen

claytreater

to Platformingunit

debutanizer

steam

to xylenesplitter

Parex unitraffinate

fuel gas

Fig. 5. Isomar process.

negligible ring destruction in the Tatoray process,there is very little hydrogen consumption. Themethyl groups are highly stable at reactionconditions and are therefore essentiallyconserved in the reaction. Most of the hydrogenconsumption can be attributed to the cracking ofthe non-aromatic impurities in the feed to theTatoray unit.

Process descriptionThe Tatoray process uses a very simple flow

scheme consisting of a fixed bed reactor and aproduct separation section (Fig. 6). The freshfeed to the Tatoray unit is combined withhydrogen-rich recycle gas, preheated byexchange with the hot reactor effluent, andvapourized in a fired heater where it is raised toreaction temperature. The hot feed vapour goesto the reactor, where it is sent down-flow over afixed bed of catalyst.

The reactor effluent is cooled by exchangewith the combined feed, mixed with make-up gasto replace the small amount of hydrogenconsumed in the reactor, and then sent to aproduct separator. Hydrogen-rich gas is taken offthe top of the separator and recycled back to thereactor. A small portion of the recycle gas ispurged to remove accumulated light ends fromthe recycle gas loop. Liquid from the bottom ofthe separator is sent to a stripper column.

The C5 overhead from the stripper is cooledand separated into gas and liquid products. Thestripper overhead gas is exported to the fuel gassystem. The overhead liquid is recycled back tothe Platforming unit debutanizer column so thatany benzene in this stream may be recovered inthe sulpholane (extraction) unit. The benzene andxylene products, together with the unreacted

toluene and A9, are taken from the bottom of thestripper and recycled back to the BTfractionation section of the aromatics complex.

Process performanceA Tatoray process unit is capable of

processing feedstocks ranging from 100 wt%toluene to 100 wt% A9�. The optimalconcentration of A9� in the feed is typically 40-60 wt%. The ability to process A9� makesmore feedstock available for xylenes productionand dramatically shifts the selectivity of the unitaway from benzene. Feeds may contain up to10% C10 aromatics.

An aromatics complex without a Tatoray unitcan produce approximately 200,000 MTA(Metric Tons per Annum) of para-xylene from25,000 BPSD (Barrels Per Stream Day) of LightArabian naphtha (160-300°F cut). If an A7Tatoray process unit (toluene feed only) is addedto the complex, the same amount of naphtha canproduce 280,000 MTA of para-xylene, a 40%increase. When an A7/A9� Tatoray process unit isadded to the complex, the endpoint of thenaphtha is increased from 300 to 340°F in order to maximize the amount of A9�

precursors in the feed. The heavier naphtha willproduce approximately 420,000 MTA para-xylene – an increase of 110% over thebase complex.

The Tatoray process produces petrochemicalgrade benzene and xylenes products. Benzenepurity with 100% toluene feed easily meets theASTM specifications for Refined 545 gradebenzene. With a feed of 50% toluene and 50% C9aromatics, the benzene product purity meets thespecifications for Refined 535 grade benzene.The xylene product from a Tatoray unit contains



heaterfuel gasfeed

surgedrum

recycle gas

toluenefrom toluene

column

C9 aromaticsfrom A9column

toluene fromParex unit

purge gasto Isomarunit

overhead liquidto Platforming unit

debutanizer

Fig. 6.Tatoray process.

an equilibrium distribution of xylene isomers andis very low in ethylbenzene. This lowethylbenzene concentration makes the xylenesproduced by the Tatoray process valuable asfeedstock to either a Parex unit or a para-xylenecrystallization unit.

PX-Plus processThe PX-Plus process selectively

disproportionates toluene to benzene andxylenes. The process is para selective, with theproduct having a para-xylene concentration in thexylene fraction of about 90%, significantly abovethe equilibrium value of 25% that is achieved bytoluene and C9 aromatic transalkylationtechnologies, such as the Tatoray process. ThePX-Plus process provides an economical way toexpand existing para-xylene facilities.

Increasing the para-xylene concentration inaromatics complex streams by adding a PX-Plusunit has significant benefits. Recoveries from asingle stage crystallizer can be increased from65% to more than 80% when the feed para-xylene concentration increases. Due to thesimilarity of operating temperature and pressureto that of many refining and petrochemicalreactor systems, existing idle equipment canoften be considered for re-use in the PX-Plusunit. The PX-Plus process can also be used forlarge-scale grassroot facilities where sufficienttoluene is available and where significantquantities of benzene are desired along withpara-xylene.

Process chemistryThe disproportionation of toluene to benzene

and xylenes proceeds via a bimolecular

intermediate. Once the bimolecular intermediatecleaves to benzene and xylene, some amount ofshifting of methyl groups can occur on the xylenemolecule. The catalyst pore structure allowsbenzene and para-xylene to escape whileinhibiting the diffusion of ortho-xylene andmeta-xylene.

Process descriptionIn the PX-Plus unit (Fig. 7), fresh toluene feed

is first combined with hydrogen-rich recycle gas,preheated by exchange with the hot reactoreffluent, and then vapourized and raised to thereaction temperature in a fired heater. Depending on the size of the unit, the reactormay be downflow or radial flow. The effluentfrom the reactor flows through the feed-effluent exchanger, is condensed, and sent tothe product (gas-liquid) separator, where recyclehydrogen is removed. The separator liquid is sentto the stripper column, where light by-productsare removed overhead. The stripper bottomsstream is then sent to benzene-toluenefractionation. High purity benzene is recoveredoverhead, and the recycle toluene is recoveredand sent back to the reactor. The para-xyleneconcentrate may then be fed directly to a single stage crystallizer, or it may be sent to theParex unit via the xylene rerun column alongwith the fresh feed mixed xylenes and the recycle isomers.

Process performanceFor a typical PX-Plus process unit,

para-xylene concentration in the product xylenesis 90%, toluene conversion per pass is 30%,benzene/xylenes mole ratio is 1.32 and benzene


AROMATICS

heaterfuel gas

feedsurgedrum

recycle gas

purge gas

overhead liquidto Platforming unit

debutanizer

steam

to B-Tfractionation

make-uphydrogen

toluene

Fig. 7.PX-Plus process.

quality is high. At 30% toluene conversion, thePX-Plus process produces less than 2 wt% lightby-products per pass. The para-xylene yieldbased on toluene converted is around 41 wt% andthe benzene yield is 46 wt%.

TAC9 processThe TAC9 process is used to selectively convert

C9-C10 aromatics into mixed xylenes. In a modernaromatics complex, the transalkylation technologies,such as the Tatoray and TAC9 processes, areintegrated between the aromatics extraction orfractionation and the xylene recovery sections of theplant. Fractionated heavy aromatics can be fed to theTAC9 unit rather than being blended into thegasoline pool or sold for solvent applications.

Incorporating transalkylation technology into anaromatics complex for the processing of toluene andC9-C10 aromatics can more than double the yield ofpara-xylene from a given naphtha feedstock. TheTAC9 process provides a more efficient means ofobtaining additional mixed xylenes from the heaviestportion of the aromatics. Processors of heavyaromatics can produce higher value products byupgrading by-product streams.

Process chemistryThe TAC9 process involves several types of

reactions of C9-C10 aromatics. Included aredisproportionation (re-arrangement of alkyl groupsbetween two identical molecules, such as toluene),transalkylation (transfer of groups between differentmolecules), and dealkylation (complete or partialremoval of an alkyl group). In the TAC9 process, thedealkylation of the alkyl groups occurs in such a waythat the methyl groups are retained.

The compounds involved are primarily alkylsubstituted aromatic compounds or alkylbenzenes.For example, the ethyl groups involved in thereactions would be those of aromatics having at leastone ethyl substitute, such as diethylbenzene,methyl-ethylbenzene (ethyl toluene), or dimethylethylbenzene (ethyl xylene). Moreover, the methylgroups involved would be those of compounds, whichare aromatics having only methyl substitutes, such astoluene, xylene, or trimethylbenzene. The highlyactive TAC9 catalyst converts almost all ethyl, propyl,and butyl groups on aromatic rings to light ends bydealkylation. The methyl groups react viadisproportionation and transalkylation. By controllingthe feed composition, a methyl balance for xyleneproduction is achieved.

The xylene yield achieved is a function of boththe methyl to phenyl (aromatic) ratio as well as theratio of ethyl groups to methyl groups in the fresh

feed. Higher ratios of methyl to ethyl groups resultin higher xylene yields. In some instances, a smallbenzene drag to an extraction unit may be requiredto optimize the yields. Mixed xylenes with very lowethylbenzene concentrations can be achieved. Thislow concentration is beneficial to operation of theIsomar and Parex units of an integrated complex.

The typical feedstock is a combination of C9 andC10� aromatics derived either from reformates orhydrotreated pyrolysis gasoline (an aromatic-rich by-product of ethylene crackers). The distribution of alkyl groups and the C9 to C10 ratio of the feedswill vary with the source as well as with theupstream pre-fractionation.

Conventional transalkylation technologies arelimited in their ability to process C10� materialprimarily due to the adverse impact that they have oncatalyst life. Higher levels result in shorter cyclesbetween catalyst regenerations, thus reducing theunit on-stream efficiencies. However, the highlyactive TAC9 catalyst converts these heaviesefficiently to higher valued products whilemaintaining an economical catalyst life.

The TAC9 reactions are conducted in a hydrogenatmosphere in order to minimize coke formation onthe catalyst. There is very low aromatic loss in theTAC9 process. Methyl groups are highly stable atreaction conditions and are therefore essentiallyconserved in the reaction. Most of the hydrogenconsumption can be attributed to the dealkylationreactions and to the cracking of the non-aromaticimpurities in the feed to the TAC9 unit.

Toray Industries introduced the currentgeneration of TAC9 catalyst in 1996. Commercialexperience has demonstrated the ability to operatefor several years without regeneration. Cycle lengthsof more than five years have been observed. Thecatalyst is regenerated using a simple carbon burnprocedure.

Process descriptionThe TAC9 process uses a very simple flow

scheme, identical to the flow scheme for the Tatorayprocess. It consists of a fixed bed reactor and aproduct separation section. The fresh feed is mixedwith recycle material then charged to the TAC9 unit.Feed material is first combined with hydrogen-richrecycle gas, preheated by exchange with the hotreactor effluent, and then vapourized in a firedheater where it is raised to reaction temperature. Thehot feed vapour then flows to the reactor, where itflows down-flow over a fixed bed of catalyst.

The reactor effluent is cooled by exchange withthe combined feed, cooled in a product condenser,and then collected in the product separator.



Hydrogen-rich gas taken off the top of the separatoris combined with make-up hydrogen and thenrecycled back to the reactor. A small portion of therecycle gas is purged from the recycle gas loop.Liquid from the bottom of the separator is sent to astabilizer column. The C5� overhead from thestabilizer is cooled and separated into gas and liquidproducts. The stabilizer overhead gas is exported tothe fuel gas system. The stabilized TAC9 product issent to product fractionation.

The process operating conditions (temperaturesand pressure) are similar to those of the Tatorayprocess. The overall conversion is close to 50%. Thiscan be tailored for the re-use of equipment in arevamp situation.

Sulpholane processThe sulphfolane process combines liquid-liquid

extraction with extractive distillation to recover highpurity aromatics from hydrocarbon mixtures, such asreformed petroleum naphtha (reformate), pyrolysisgasoline, or Coke Oven Light Oil (COLO).Contaminants that are the most difficult to eliminatein the extraction section are easiest to eliminate inthe extractive distillation section and vice versa. Thishybrid combination of techniques allows sulpholaneunits to process feedstocks of much broader boilingrange than would be possible by either techniquealone. A single sulpholane unit can be used forsimultaneous recovery of high-purity C6-C9aromatics, with individual aromatic componentsrecovered downstream by simple fractionation.Typically, when just benzene or toluene isproduced, the sulpholane unit can be built as anExtractive Distillation (ED) unit only and the extractor can be eliminated, thereby simplifying the design.

The sulpholane process takes its name from thesolvent used: tetrahydrothiophene 1, 1-dioxide, or‘sulpholane’. Sulpholane is the most efficientsolvent available for the recovery of aromatics. Mostextraction units can be made to operate at highpurity and recovery by circulating more and moresolvent. Because the sulpholane solvent exhibitshigher selectivity and capacity for aromatics thanany other commercial extraction solvent, sulpholaneunits operate at the lowest available solvent-to-feedratio for any given reformate feedstock. Therefore,for reformate applications, a sulpholane unit is lessexpensive to build and operate than any other type ofextraction unit.

A sulpholane unit is usually incorporated withinan aromatics complex to recover high-purity benzeneand toluene products from reformate. In a modern,fully integrated aromatics complex, the sulpholane

unit is located downstream of the reformate splitter column. The C6-C7 fraction from theoverhead of the reformate splitter is fed to thesulpholane unit. The aromatic extract from thesulpholane unit is clay treated to remove traceolefins, and individual benzene and tolueneproducts are recovered by simple fractionation. Theraffinate from the sulpholane unit is usually blendedinto the gasoline pool or used in aliphatic solvents.Benzene must always be recovered by extraction orextractive distillation in order to meet purityspecifications for petrochemical applications.Toluene must be extracted for direct use inpetrochemical applications and is usually extractedbefore being fed to a dealkylation ordisproportionation unit for production of additionalbenzene and xylenes. Modern CCR Platformingunits operate at such high severity that the C8�

fraction of the reformate does not contain anysignificant amount of non-aromatic impurities andmay be sent directly to the xylenes recovery sectionof the plant without extraction. However, the C8�

fraction of pygas and COLO streams containssignificant amounts of non-aromatic impurities and,therefore, must be extracted before either beingrecovered as marketable mixed xylenes or sent toxylene recovery.

Process descriptionAs shown in Fig. 8, fresh feed enters the extractor

and flows upward, countercurrent to a stream of leansolvent. As the feed flows through the extractor,aromatics are selectively dissolved in the solvent. Araffinate stream, very low in aromatics content, iswithdrawn from the top of the extractor. The richsolvent, loaded with aromatics, exits the bottom of theextractor and enters the stripper. The non-aromaticcomponents having volatilities higher than that ofbenzene are completely separated from the solvent byextractive distillation and removed overhead alongwith a small quantity of aromatics. This overheadstream is recycled to the extractor where the light non-aromatics displace the heavy non-aromatics fromthe solvent phase leaving the bottom of the extractor.The bottoms stream from the stripper, substantiallyfree of non-aromatic impurities, is sent to the recovery column, where the aromatic product isseparated from the solvent.

Because of the large difference in boilingpoint between the sulpholane solvent and theheaviest aromatic component, this separation isaccomplished easily, with minimal energy input.To minimize solvent temperatures, the recoverycolumn is operated under vacuum. Lean solventfrom the bottom of the recovery column is


AROMATICS

returned to the extractor. The extract is recoveredoverhead and sent on to distillation columnsdownstream for recovery of the individualbenzene and toluene products. The raffinatestream exits the top of the extractor and isdirected to the raffinate wash column. In thewash column, the raffinate is contacted withwater to remove dissolved solvent. The solvent-rich water is vapourized in the waterstripper by exchange with hot circulating solventand then used as stripping steam in the recoverycolumn. Accumulated solvent from the bottom ofthe water stripper is pumped back to the recoverycolumn. The raffinate product exits the top of theraffinate wash column. The amount of sulpholanesolvent retained in the raffinate is negligible. Theraffinate product is commonly used for gasolineblending or aliphatic solvent applications. Undernormal operating conditions, the sulpholanesolvent undergoes only minor oxidativedegradation. A small solvent regenerator isincluded in the design of the unit as a safeguardagainst the possibility of air leaking into the unit. During normal operation, a small slip-stream of circulating solvent is directed tothe solvent regenerator for removal of oxidizedsolvent.

The extract product from a sulpholane unitmay contain trace amounts of olefins and otherimpurities which would adversely affect the acidwash colour tests of the final benzene andtoluene products. To eliminate these traceimpurities, the extract is clay-treated prior to

fractionation. Clay treating is done at very mildconditions and clay consumption is minimal.

The treated extract is directed to thearomatics fractionation section where high-puritybenzene, toluene, and sometimes mixed xylenesare recovered. The design of the aromaticsfractionation section varies depending on theparticular processing requirements of thecustomer.

Process performanceThe performance of the sulpholane process has

been well demonstrated in more than 130 operatingunits. The recovery of benzene exceeds 99.9 wt%, andrecovery of toluene is typically 99.8 wt%. Thesulpholane process is also efficient at recovery ofheavier aromatics if necessary. Typical recovery ofxylenes exceeds 98 wt%, and 99 wt% recovery hasbeen demonstrated commercially with rich feedstocks.

Sulpholane units routinely produce a benzeneproduct with a 5.5°C solidification point or better, andmany commercial units produce benzene containingless than 100 ppm non-aromatic impurities. Thetoluene and C8 aromatics products from a sulpholaneunit are also of extremely high purity, easily exceedingnitration grade specifications. In fact, the ultimatepurity of all of the aromatic products is usually moredependent on the design and proper operation of thedownstream fractionation section than on theextraction efficiency of the sulpholane unit itself. The purity and recovery performance of an aromaticsextraction unit is largely a function of energyconsumption.



feedsurgedrum

raffinateproduct

to storage

solv

ent

rege

nera

tor

to ejector

extract to claytreaters in BTfractionationunit

recoverycolumn

steam

steam

steam

Fig. 8. Sulpholaneprocess.

In general, higher solvent circulation rates result inbetter performance, but at the expense of higherenergy consumption. The sulpholane processdemonstrates the lowest energy consumption of anycommercial aromatics extraction technology. A typicalsulpholane unit consumes 275-300 kcal of energy perkilogram of extract produced, even when operating at99.99 wt% benzene purity and 99.95 wt% recovery.Sulpholane units are also designed to efficientlyrecover solvent for recycle within the unit. Expectedsolution losses of sulpholane solvent are less than 5ppm of the fresh feed rate to the unit.

Bibliography

Gary J.H., Handwerk G.E. (1984) Petroleum refining.Technology and economics, New York-Basel, MarcelDekker.

Johnson J.A. (1986) Aromatics complexes, in: Meyers R.A.(editor in chief) Handbook of petroleum refining processes,New York-London, McGraw-Hill.

Vladas ZukauskasCopyright 2004 UOP LLC

All rights reservedUsed with permission


AROMATICS


10.6.2 Aromatic intermediates in petrochemical industry

IntroductionThe alkylation reaction of aromatic compounds is

widely used in the chemical industry for the productionof important intermediates such as ethyl benzene andcumene. Fig. 1 shows the main alkylations employed inindustry for the transformation of benzene, toluene andxylene, by reaction with olefins, into isopropyl benzene(cumene), diisopropyl benzene, ethyl benzene, diethylbenzene, C10-C14 Linear Alkyl Benzenes (LAB),cymene, isobutyl benzene, and o-tolyl pentene. In Fig. 1the principal derivatives of these intermediates are alsoindicated. The alkylations shown are of two types: thealkylation of the aromatic ring catalysed by acids andthat of the side chain, catalysed by bases.

From the point of view of production volumes, theacid-catalysed alkylations of benzene are the mostimportant: in 2004 about 75% of the benzeneproduced in the world (36.5 million tons) wasalkylated with olefins, 71% of which was to produceethyl benzene and cumene (Table 1).

Alkylation of aromatic hydrocarbons with olefinsThe alkylation reaction of aromatic hydrocarbons

can be performed using various alkylating agents:

alkyl halides, alcohols, alkyl sulphates, and olefins.Olefins are the most extensively used alkylating agentsin the petrochemical industry. The alkylation reactionof the aromatic ring with olefins is an exothermicreaction and thus is favoured, from the thermodynamicpoint of view, at low temperatures. In fact theequilibrium constant diminishes with increasingtemperature, as illustrated in Fig. 2 for the alkylation ofbenzene with ethylene and propylene. The enthalpy ofreaction at 25°C in the gaseous state is –105.51 and–99.65 kJ/mol, for the formation of ethyl benzene andcumene respectively. Thus the formation of these alkylaromatics is accompanied by a release of energy, in theform of heat, which needs to be taken into account inthe design of production plants.

Acid catalysisDifferent types of acids are used as catalysts for the

alkylation of aromatic hydrocarbons: a) metal halides,such as aluminium and gallium chloride and boriumfluoride; b) mixed oxides and zeolites; c) protonic acidssuch as sulphuric acid, hydrofluoric acid and phosphoricacid; and d) sulphonic resins. The most active ones arethe Brönsted acids, which contain an acidic proton. Themetal halides, which are Lewis acids, are not very activealkylation catalysts if used as such and must beactivated by means of adding small quantitites of a

� �

OH

SO3NaH3CCHCH3

CH�CH2 C2H5

C2H5C2H5

H3CCHCH3

H3CCHCH3

H3CH3C

CH3CH3

CH3

CH3CH3

CH3

OH

OH

CH3

CH2

C

H

ibuprofen

CH3

OH

Fig. 1. Main aromatichydrocarbon alkylations witholefins of industrial interest.


co-catalyst, such as a hydrohalic acid. The co-catalystreacts with the Lewis acid thus generating a Brönstedacid. These catalysts are also known as Friedel-Craftscatalysts and are still extensively used in alkylation,even if the new processes use solid acid catalysts. TheFriedel-Crafts catalysts and mineral acids in general aretoxic and highly corrosive and therefore difficult tohandle and store, as they corrode the containers, thepipes, and apparatus in which they are used. At the endof the reaction they are mixed with the product andmust be separated by means of operations which areexpensive and energy consuming. It is often necessaryto neutralize them with bases and the salts obtained inthis way are then separated by washing with water. Inthis manner the acid catalyst is not recovered and thesalt water produced must be disposed of, resulting in

problems related to the environmental impact. For all ofthese reasons the solid acid catalysts, and in particularthe zeolites, are preferred in the new generation oftechnologies (Perego and Ingallina, 2004).

Zeolites are high porosity crystalline aluminiumsilicates, exhibiting regular pores with diameters rangingfrom 0.3 to 1 nm. Depending on the atomic structure ofany given zeolite, up to 50% of its volume can beformed by porous cavities. The fundamental unit of thezeolitic structure is the tetrahedron formed by Si�4 andAl�3 attached to four atoms of oxygen. The tetrahedraare connected together by sharing an oxygen atom toform polymer chains. The formation of the threedimensional structure is due to the fact that all fouratoms of oxygen in each tetrahedron are shared withother tetrahedra. On the basis of the number oftetrahedra that make up the opening of the zeoliticchannels, the zeolites are classified in the following way:small pores (8 tetrahedra), medium pores (10 tetrahedra),large pores (12 tetrahedra) and extra large pores (>12tetrahedra). Another characteristic of the zeoliticporosity is the presence of interconnections between thenetworks of channels, which can be one-dimensional,two-dimensional or three-dimensional. The presence ofaluminium in tetrahedral coordination generates negativecharges that are locally neutralized by the cations presentin the zeolitic channels; by exchanging these cationswith the proton a zeolitic acid is obtained.

Zeolitic catalysis is characterized by shapeselectivity (Csicsery, 1995). This principle is quitesimple: the system of pores of a zeolite can regulatethe entrance of the molecules of the reagents anddetermine the dimensions of the intermediates and theproducts. The zeolites of interest for the alkylation ofaromatics are predominantly of medium and large poresize, as shown in Table 2, together with the mostimportant structural characteristics.

Reaction mechanism. The alkylation mechanism firstinvolves the formation of an electrophile E� by means ofthe interaction between the olefin and the acid; followedby electrophilic attack on the aromatic ring (Ar�H),with the formation of an intermediate [E�Ar�H]�,known also as the Wheland intermediate or arenium ion.From this intermediate, an alkyl aromatic is formed byelimination of H�, i.e. of a proton. In the case ofalkylation of benzene with ethylene, catalysed by a HAacid, the reaction sequence:


AROMATICS

�4

0

4

3.5 2.5

1,000/T (K)

1.5 0.5

8

12

log

K

ethyl benzenecumene

Fig. 2. Dependence of the equilibrium constants of thealkylation reactions of benzene with ethylene and propyleneto produce ethyl benzene and cumene.

�

CH2 � CH2�HA

CH2 CH3H

A� � H�A�

CH3CH2�A�

CH2CH3

�

Derivative Millions tons of benzene %

Ethyl benzene 19.5 53.4

Cumene 6.3 17.3

LAB 1.3 3.5

Cyclohexane 5.0 13.8

Nitrobenzene 2.3 6.3

Other 2.1 5.7

Total 36.5 100

Table 1. Breakdown of benzene usein the chemical industry

The activated species CH3�CH2�A� is often

represented as a free carbocation CH3�CH2�. This

is, however, a qualitative representation, and in factthe completely free carbocation is never obtained:the protonation of ethylene produces a singleprimary carbocation. In the case of higher olefinsthe protonation can form two carbocations; forexample from propylene the secondary carbocation(i-propyl, CH3CH�CH3) and the primary one, (n-propyl, CH3CH2 CH2

�), may be generated. Therelative stability of the carbocations increases inthe order primary<secondary<tertiary andinfluences the rate and selectivity of the alkylationreaction.

Kinetics. The effect of the nature of the alkyl groupon the reaction rate is shown in Table 3. Theisopropylation of toluene is approximately 1,460 timesfaster than its ethylation, using GaBr3 as the catalystand the alkyl bromide as the alkylating agent. Similardata were reported for the alkylation of benzene witholefins, with a zeolite catalyst (REY, rare earthexchanged Y zeolite): at 100°C propylation is about300 times faster than ethylation (Beck and Haag,1997).

These data agree with the protonic affinity (Vogel,1985), i.e. the tendency of an olefin to be protonatedthus creating the corresponding carbocation:

ethylene to ethyl 672 kJ/mol propylene to i-propyl 755 kJ/mol i-butene to t-butyl 810 kJ/mol

Selectivity. The products of alkylations using acidcatalysts are those derived from the most stablecarbocations. For example, in the case of the alkylationwith propylene, the i-propyl benzene (cumene) ispractically the only product of mono-alkylation,because the formation of the i-propyl carbocation isfavoured with respect to the n-propyl carbocation: thedifference in the enthalpies of formation between thetwo carbocations in the gas phase is 67 kJ/mol. Inaddition to the alkylation of the aromatic ring, the olefincan also follow other reaction paths: it can react withitself forming higher oligomers, as in the case of theproduction of cumene accompanied by the formation ofpropylene oligomers (for example nonene), or it canisomerize forming other olefins that, in their turn, canproduce other alkylation products, such as in the case ofthe alkylation of benzene with 1-dodecene where,together with 2-phenyl dodecane, other isomers are alsoformed (3-,4-,5-,6-phenyl dodecane).

Oligomerization and isomerization are parallelreactions to alkylation and reduce its selectivity. Both arereactions which are catalysed by acids and can be limitedthrough the use of very selective alkylation catalysts orby means of suitable operating conditions. Theselectivity is also affected by consecutive reactions. Afterthe first alkylation the aromatic substrate can undergosuccessive alkylations, forming polyalkylatedby-products. The presence of alkyl substituents on thearomatic ring increases its reactivity, due to their abilityto favour the delocalization of the positive charge on theWheland intermediate. In fact, the presence of alkylsubstituents such as ethyl or i-propyl on the aromaticring increases the rate of Friedel-Crafts alkylation by afactor of 1.4 to 3.2 with respect to unsubstituted benzene



ZeoliteIZA code Dimensional

Pores Dimensions of(International Zeolite structure of pores (nm)Association) channels

Beta BEA 3D large0.66�0.670.56�0.56

Mordenite MOR 1D large 0.65�0.70

Y FAU 3D large 0.74�0.74

MCM-22 MWW 3D medium0.55�0.400.51�0.41

ZSM-5 MFI 3D medium0.53�0.560.51�0.55

Table 2. Relevant zeolites in catalytic alkylation

Alkyl group Relative rate

Methyl 1.0

Ethyl 13.7

i-propyl 20,000

Table 3. Relative rate of the alkylation of toluene(Allen and Yats, 1961)

(Olah, 1973). Various methods are employed in order tomaximize the yield of the monoalkylate. The mostobvious is to operate with a large molar excess, or with ahigh aromatic/olefin ratio. This also makes it possible tolimit the sub-production of oligomers (Norris Shreveand Albright, 1958). The disadvantage is constituted bythe necessity of separating the aromatic excess andrecycling it, with elevated energy costs. Anotherstratagem involves the use of very selective catalystssuch as zeolites, capable of reducing the formation of thebigger polyalkylates and to favour the formation of themonoalkylates. For example, in the alkylation ofbenzene with propylene catalysed by Beta zeolite:

the ratio between the rate constants for the formationof the products di- and mono-alkylates is equal to 0.54(Perego et al., 1999).

Alkylation of benzene

Ethyl benzeneIn 2004 the world production of Ethyl Benzene

(EB) was about 26 million tons, with demand growing

on average at a rate of 4-5% per year. Almost all EB isused for the production of styrene, a raw material forthermoplastic polymers and elastomers.

Alkylation of benzene to EB. The main reactionsthat take place during the alkylation of benzene withethylene, catalysed by acids, are summarized in Fig. 3.

The first stage is represented by the formation of theethyl carbocation, which then follows one of twoprincipal reaction paths. Either it reacts with benzene togive EB, following which, by successive alkylations,forms Di-Ethyl and Tri-Ethyl Benzene (DEB and TEB),or it reacts with an other molecule of ethylene to form aC4 carbocation, that can then successively undergoalkylation, oligomerization, isomerization and crackinggiving other alkyl benzenes and olefins. To a verylimited extent EB can alkylate benzene to 1,1-diphenylethane. It is important to emphasize that DEB and TEBcan react with benzene to give EB. This reaction, knownas transalkylation, is an equilibrium reaction:

and under suitable conditions can already take placeduring alkylation. The traditional process for theproduction of EB was developed around 1930. Thecatalyst used was AlCl3-HCl and all the operationswere carried out in an agitated reactor, undersomewhat gentle conditions: 170°C and 0.7 MPa. Onleaving the reactor, after separation, the polyethyl


AROMATICS

� CH2�CH�CH3

CH3�CH�CH3

k1

CH2�CH�CH3

CH3�CH�CH3

CH3�CH�CH3

k2

CH2 = CH2 CH3�CH2�

CH2CH3

CH2CH3C2H4

CHCH3

C2, C4, C5, C7 olefins

other alkylates

higher alkylbenzenes


C2H4

C4H9�

C6H13

C2 H

5�

H�

�H�

�H

�

�H�

cracking

�H�

�H�

�H�

C4 olefins

C6 olefins

CH2CH3

�

Fig. 3. Main reactionsduring the alkylation ofbenzene with ethylene.

C2H5 C2H5

C2H5

� 2

benzenes (mainly DEB and TEB) were recycled in thealkylation reactor where they were converted, in thepresence of an excess of benzene, by transalkylation,reaching a conversion near to the thermodynamicequilibrium. The equilibrium composition is a functionof the ratio of ethylene/benzene; this ratio is typicallyin the range between 0.35 and 0.55 (Franck andStadelhofer, 1988).

In order to overcome the problems connectedwith the use of AlCl3, starting in the mid 1960s,various zeolitic catalysts were tested for thisreaction. In 1976 Mobil-Badger started up the firstindustrial plant for the production of EB in the gasphase, with a fixed-bed reactor, loaded with acatalyst based on ZSM-5. The reactor operatedunder high temperature (390-450°C) and pressure(1.5-2 MPa) conditions. As in the process withAlCl3, after separation, the polyalkylates wererecycled to the reactor for transalkylation. Due tothe deactivation related to the deposition of carbonresidues (coke) in the zeolitic pores, the catalysthad to be regenerated every 40-60 days. Theregeneration was performed in situ, by blowing inair to allow the combustion of the coke. The highfrequency of this operation made it necessary tohave two reactors, one for the regeneration and onefor the reaction, in order to guarantee continuousproduction. This process, used commercially from1980 onwards, was successively improved by theaddition of a reactor dedicated to thetransalkylation of the polyethyl benzenes, thusobtaining an improvement both in the yield and inthe catalyst life (Wang,1993).

A considerable improvement came later, obtainedby UOP/Lummus/Unocal with the development of aprocess in the liquid phase. The advantage of theliquid phase is represented by better thermal controlthat is reflected in an extension of the catalyst life. Inthis way the regenerations are less frequent and can beconducted on the catalyst which is discharged from thereactor and put into dedicated ovens. Due to problemsrelated to diffusion control, medium pore zeolites suchas ZSM-5, were not suitable for use in the liquidphase. For this reason, in the new process a large-porezeolite, called Y zeolite, was used. The process wascommercially used for the first time in Japan in 1990(Narsolis et al., 1997).

Other large-pore zeolites have been shown to besuitable for liquid phase alkylation of benzene withethylene (for example, L, Omega, ZSM-12, Beta).In particular the zeolite Beta has turned out to bemore selective than an Ultra-Stabilised Y zeolite(USY), with a global selectivity (EB�DEB�TEB)of 99.3% against 91.1% for zeolite Y. Both zeoliteshave a three-dimensional system of channels, but

the presence of large cavities (1.2 nm in diameter) atthe intersections of the channels (in zeolite Y) isprobably the cause of the formation of a largeamount of by-products that, apart from reducing theselectivity, produce a faster deactivation of thecatalyst (Bellussi et al., 1995).

Very interesting results have also been obtainedwith MCM-22, a zeolite of medium pore size,characterized by two systems of channels independentof each other, one of which containing large cavitiesopen to the exterior on the external surface, withopenings formed by 12 tetrahedra and withdimensions of 0.71�0.71 nm. Thanks to thispeculiarity, MCM-22 demonstrates a catalytic activitycomparable to USY, but inferior to that of Beta.However, the selectivity is higher, both with respect toUSY and to Beta, in as much that the formation ofDEB and TEB, is particularly reduced (Cheng etal.,1999). MCM-22 was applied to a liquid phaseprocess called EBMax which has beencommercialized by Exxon/Mobil since 1995.

The zeolite Beta is the process catalystdeveloped by Polimeri Europa for the productionof EB. After some years of evaluation in a pilotplant, the catalyst was evaluated in an industrialreactor of an existing EB plant, starting in 2001.The catalyst, based on zeolite Beta, demonstratedexceptional performance both from the point ofview of the consumption of raw materials and ofthe quality of the EB produced. The flow diagramof the Polimeri Europa process is shown in Fig. 4(Girotti et al., 2004). The catalyst is distributed inthe reactor on more beds and the supply ofethylene is directed over these in such a way as tocreate a much higher local ratio ofbenzene/ethylene with respect to the global ratio,for the reasons described above. This is a stratagemwhich is used in all alkylation processes usingreactors with catalytic beds.

Since 1990 UOP/Lummus has improved its ownliquid phase process (Narsolis et al., 1997), whichnow is used commercially under the name EBOne.The catalyst, originally a zeolite Y, now consists of amodified zeolite Beta. The improvements made tothe catalyst have also been extended to the processdeveloped by CDTECH (a consortium betweenABB Lummus and Chemical Research andLicensing), based on catalytic distillation thatcombines the reaction and distillation into a singleoperation. The catalyst, packaged in ‘bales’, ispositioned on the plates of the distillation column(Fig. 5). On each individual plate the ratio ofbenzene/ethylene proves to be very high (�1,000),due to the resistance to mass transfer and to theliquid-vapour equilibrium, with the resulting



advantages in terms of selectivity. In addition, theheat of reaction is exploited to distil the benzene,thus making an energy saving. This process, knownas CDTECH EB, is particularly suitable for dilutedflows of ethylene (Cho and Zhu, 2003). Of over 70industrial EB plants in the world, in 2002, only 24%still used AlCl3-HCl as the catalyst; the others wereusing zeolitic catalysts: 40% in gas phase and 36%in the liquid phase (Perego and Ingallina, 2002).

Isopropyl benzene (cumene)World production of cumene in 2004 was 9.5

million tons; it is almost exclusively used for theproduction of acetone and phenol, and a growth of 5%per year is predicted for the consumption of phenol.

Alkylation of benzene with propylene. Thealkylation reaction of benzene with propylene is verysimilar to that with ethylene (Fig. 6).

The i-propyl carbocation reacts with benzene togive cumene and by successive alkylations to givedi- and tri-isopropyl benzene. Di- and tri-isopropylbenzene can transalkylate to cumene in thepresence of an excess of benzene. The carbocation

can, in addition, react with propylene producing C6carbocations which evolve by means ofoligomerization, cracking and alkylation, to givehigher oligomers and other alkyl benzenes. Smallquantities of n-propyl benzene are also obtained byisomerization of cumene. This represents a verycritical aspect, as n-propyl benzene cannot beseparated by simple distillation and thus itsformation affects the final quality of the cumene.

The demand for cumene as a high-octane additivefor military aeroplanes in the Second World War led tothe development of the first process based on the useof sulphuric acid. The problems related to the use of afree acid were overcome in the 1940s, with theintroduction, by UOP, of a catalyst based on SupportedPhosphoric Acid (SPA). This technology is still widelyused today: the catalyst is loaded in a fixed-bed reactoroperating in the liquid phase (180-240°C; 3-4 MPa).The formation of polyalkylates, that are nottransalkylated by SPA, and the oligomers of propylene,is minimized by operating with a high ratio ofbenzene/propylene (from 5 to 10). Also in this case,the ratio is further increased by subdividing the supply


AROMATICS

claytreatment

ethylenecompression

ethylene

benz

ene

colu

mn

alkylationsection

transalkylationsection

drag

ben

zene

colu

mn

colu

mn

EB

colu

mn

PEB

benzene

purge

benzene

oily water

EB

to vacuumsystem

flux oil

PEB

Fig. 4. Flow scheme for the liquid phase process for the production of EB.

of propylene to individual catalytic beds. The SPAcatalyst, even though supported, nevertheless generatesproblems of corrosion, due to a release of free acid; inaddition, at the end of its life cycle it can not beregenerated.

In the 1970s, Monsanto-Lummus introduced a newtechnology based on the use of AlCl3-HCl, verysimilar to that used for EB. The advantage of AlCl3resides in its ability to catalyse the transalkylation ofthe polyalkylates, that can then be recycled to thealkylation reactor, in contrast to the SPA catalyst.However, only a few plants have been implementedwith this technology.

The search for zeolitic catalysts for the productionof cumene has been in many ways similar to that forEB, even if many more years were necesssary to arriveat a significant result, than with EB. This is due mainlyto the fact that the zeolite ZSM-5, the catalyst forMobil’s EB process, demonstrated a major limitationin the cumene reaction, represented by the elevatedlevel of co-production of n-propyl benzene. On theother hand, ZSM-5, being of medium pore size, is notsufficiently active in the liquid phase (Bellussi et al.,1995). Also for cumene, a noticeable improvementwas thus obtained by operating in the liquid phasewith large pore zeolites. With catalysts based on thesezeolites, around the middle of the 1990s new processeswere publicised and the first industrial demonstrationsdeveloped by Dow-Kellogg, Mobil-Raytheon,CDTech, EniChem and UOP. In all cases, it was amatter of improvements made in existing plants,through the substitution of SPA with a zeolitic catalyst.



lights

benzene

catalytic distillation section

regular distillation section

ethylene

ethyl benzene and heavier

Fig. 5. Catalytic distillation conceptfor the production of EB.

CH2 = CH�CH3 CH3�CH��CH3

CH3

CH3CH2�CH2�CH3C3H6

C2, C4, C5, C7 olefins



C3H6

C6H13

C9H19

C3H7�

CH3

CH

CHCH3

H�

�H�

�H�

�H�

cracking

�H�

�H�

�H�

C6 olefins

C6 olefins

CH

CH3 CH3

other alkylates

�

�

Fig. 6. The alkylationreaction of benzene with propylene.


AROMATICS

In Fig. 7 the type of reactor and the zeolites used inthese processes are shown.

Starting in March 1996, Enichem (now PolimeriEuropa) evaluated a new catalyst based on zeolite Betaon a reactor of the plant for cumene production at

Porto Torres. The results were very positive both withregard to the specific consumption of raw materialsand to the quality of the cumene. Fig. 8 compares thequality of cumene obtained with a zeolite catalyst andwith an SPA catalyst, in terms of the concentration ofimpurities. On the basis of these results the wholeplant was converted to zeolite Beta. After years ofoperation the new catalyst has shown a high level ofstabilty, reaching a production of more than 30,000tons of cumene per ton of catalyst, compared with the1,500 tons obtained per ton of SPA. The flow diagramof the Polimeri Europa cumene production process isshown in Fig. 9 (Girotti et al., 2004). In 2001, out ofaround 40 cumene plants existing in the world, 14were already operating with a zeolitic catalyst (Degnanet al., 2001).

References

Allen R.H., Yats L.D. (1961) Kinetics of three compoundequilibrations. V: Concurrent alkylation and isomerization,«Journal of the American Chemical Society», 83, 2799-2805.

Beck J.S., Haag W.O. (1997) Alkylation of aromatics, in: G.Ertl et al. (edited by) Handbook of heterogeneous catalysis,Weinheim, VCH, 2123-2139.

Bellussi G. et al. (1995) Liquid-phase alkylation of benzenewith light olefins catalyzed by b zeolites, «Journal ofCatalysis», 157, 227-234.

Cheng J. C. et al. (1999) A comparison of zeolites MCM-22,Beta and USY for liquid phase alkylation of benzene with

company ExxonMobil UOP CD-tech Dow-Kellog Polimeri Europa

zeolite MCM-22 Beta Y Mordenite Beta

reactorrecycledfixed bed

fixed bedcatalytic

distillationfixed bed fixed bed

process Q-max CDCumene 3-DDM

Fig. 7. Processes for the production of cumene with zeolitic catalysts.

zeolite

SPA

800

ppm

700

600

500

400

300

200

100

non aromatics n-propylbenzene

t-butyl benzene bromo index

a-methylstyrene0

Fig. 8. Comparison of catalyst performance between a zeoliteBeta catalyst and an SPA catalyst.



ethylene, in: Studies in surface science and catalysis, 121,Amsterdam, Elsevier, 53-60.

Cho S., Zhu W. (2003) Ethylbenzene/styrene monomertechnological advancements enable dramatic improvementsin unit capacities and project economics, in: Proceedingsof the ERTC petrochemical conference, Paris (France), 3-5March.

Csicsery S.M. (1995) The future of shape selective catalysis,in: Studies in surface science and catalysis, 94, Amsterdam,Elsevier, 1-12.

Degnan T.F. et al. (2001) Alkylation of aromatics withethylene and propylene. Recent developments incommercial processes, «Applied Catalysis A. General»,221, 283-294.

Franck H.G., Stadelhofer J.W. (1988) Industrial aromaticchemistry, Berlin, Springer, 134.

Girotti G. et al. (2004) Zeolite catalysts the way forward,«Hydrocarbon Engineering», November.

Narsolis F. et al. (1997) High performance catalyst for liquidphase EB technology, «Petroleum Technology Quarterly»,Summer, 77-81.

Norris Shreve R., Albright L.F. (1958) Alkylation, in:

Groggins P.H. (editor in chief) Unit processes in organicsynthesis, New York, McGraw-Hill, 804-855.

Olah G.A. (1973) Friedel-crafts chemistry, New York, JohnWiley, 35-36.

Perego C., Ingallina P. (2002) Recent advances in theindustrial alkylation of aromatics. New catalysts and newprocesses, «Catalysis Today» , 73, 3-22.

Perego C., Ingallina P. (2004) The combining alkylationand transalkylation for alkylaromatic production, «GreenChemistry», 6, 274-279.

Perego C. et al. (1999) Development and industrial applicationof a new b zeolite catalyst for the production of benzene,in: Proceedings of 12th International zeolite conference,Warrendale (PA), Material Reasearch Society, 4v.; v. I, 575.

Vogel P. (1985) Carbocation chemistry, Amsterdam, Elsevier, 74.Wang S.-H. (1993) Styrene, Process Economics Program

Report 33C, Supplement C, SRI Consulting.

Carlo PeregoEniTecnologie

San Donato Milanese, Milano, Italy

benz

ene

colu

mn

cum

ene

colu

mn

diis

opro

pylb

enze

neco

lum

n

zeolitecatalyst

depr

opan

izerzeolite catalyst

drag

ben

zene

colu

mn

purgefreshbenzene

fresh propyleneal

kyla

tion

reac

tor

tran

salk

ylat

ion

reac

tor

polyalkylbenzenes

recycle benzene

claytreatment

oily water

pure cumene

heavier to fuel

LPG

Fig. 9. Flow scheme of the liquid phase process for the production of cumene.

Aromatics Và Các Phương Pháp Sản Xuất

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